Destructive hydrogenation of hydrocarbon mixtures containing difficultly vaporizablecomponents



Jan 18, 1955 v. J. ANHoRN ETAL- 2,700,014

DEsTRucTrvE HYDROGENATION oF HYDRocARBoN MIXTURES coNTAmmG-DIFFICULTLY vApoRIzABLE COMPONENTS Filed May 51, 195o PRODI.: @its To REOOVQFCF IN VEN TORS United States Patent() DESTRUCTIVE HYDROGENATION OF HYDRO- CARBON MIXTURES CONTAINING DIFFI- CULTLY VAPORIZABLE COMPONENTS Victor J. Anhorn and Charles W. Montgomery, Oakmont,

Pa., assignors to Gulf Research & Development Company, Pittsburgh, Pa., a corporation of Delaware Application May 31, 1950, Serial No. 165,208

e 9 Claims. (Cl. 196-53) This invention relates to an improved process for the destructive hydrogenation of hydrocarbon mixtures which contain diicultly vaporizable hydrocarbon components.

The destructive hydrogenation of hydrocarbon mixtures which contain difcultly vaporizable components has in the past been accomplished primarily by contacting the hydrocarbon mixture, while in the liquid phase, with hydrogen and with a hydrogenation catalyst under destructive hydrogenation conditions. The hydrogenation catalyst was either suspended in the liquid or the hydrocarbon liquid was contacted with a stationary bed of the catalyst. Such operations have been unsatisfactory, mainly because they entailed the use of excessively high hydrogen pressures, usually of Iabout 3000 or more pounds per square inch, to avoid coke formation. These highpressures entail high operating and equipment costs. Also large amounts of hydrogen, which is `an expensive material, are consumed. This is largely due to the fact that high pressures cause excessive hydrogenation of aromatic compounds to take place. Due to this hydrogenation, the gasoline products are of relatively poor quality. It is Ialso common practice to convert stocks containing difficultly vaporizable components by irst treating these diicultly vaporizable components in a ook-ing still and then cracking the heavier portion of the liquid products so produced. This procedure is uneconomical because of the poor yields which result from excessive coke and gas formation. It has also been suggested to directly introduce a liquid stock into a fluidized bed of cracking catalyst as in conventional fluid catalytic cracking. This procedure also is uneconomical because of excessive loss due to coke and gas formation and because of excessive catalyst regeneration requirements. Therefore, no single procedure heretofore known resulted in conversion of such stocks to `a good quality product, in good yield, with relatively low carbon and gas format-ion and low hydrogen consumption.

This invention has for its object to provide a process for economically converting hydrocarbon mixtures which contain difiicultly vaporizable, or so-called residuaL components into lower boiling hydrocarbons of improved quality in good yields. Another object is to provide an improved process for the destructive hydrogen-ation of hydrocarbon mixtures containing components which are diflicultly vaporizable. Another object is to provide an improved process for treating hydrocarbons containingk diflicultly vaporizable components whereby a high degree of conversion to lower boiling liquid products is obtained. A still further object is to provide an improved process for the simultaneous destructive hydrogenation `and desulfurization of hydrocarbon mix-tures containing diflicultly vaporizable components and sulfur compounds. Another object is to provide a process for destructively hydrogenating a hydrocarbon mixture containing components which are diiicultly vaporizable, which process will give a good quality product, in good yield, with relatively low gas and carbon formation and with relatively low hydrogen consumption. Other objects will appear hereinafter.

These and other objects are accomplished by our invention, which includes the catalytic conversion to lower boiling point products of a hydrocarbon mixture which contains diicultly vaporizable components by contacting this hydrocarbon mixture, while it contains a substantial amount of the difficultly vaporiz-able components in liquid phase, with hydrogen and particles of a hydrogenaton catalyst, at a temperature between about 750 and 950 F.

ICC

and at a pressure between about 250 and 2000 p. s. i. g., while maintaining a substantial amount of the liquid portion of the hydrocarbon mixture absorbed on the catalyst particles and while main-taining the catalyst particles Y -in a turbulent and suspended condition during said contacting. The catalyst particles Iare then separated from the hydrogen .and the hydrocarbon reaction products. We have found that when operating in this manner the liquid portion of the hydrocarbon mixture is absorbed on the catalyst particles where it becomes converted into lower boiling hydrocarbons and especially into high yields of gasoline of good quality, this being accomplished without causing agglomeration of the catalyst particles and without resulting in the formation of a large amount of gas land coke. v The vapor portion of the charge is also very effectively converted to give a high yield of desirable liquid products. Due to the relatively low pressures employed, hydrogen consumption and cost of equipment is considerably reduced.

iIn the following examples and description we have `set forth several of the preferred embodiments of our invention but it is to be understood that they are given by Way of illustration and not in limitation thereof.

We have illustrated in the accompanying drawing a diagrammatic elevation, partly in section, of apparatus in which our invention can be carried out.

Referring to the drawing, numerals 1 and 2 designate reactors. which are alternately on-stream and alternately regenerated. Reactors 1 and 2 are partiallytlilled with a body of powdered hydrogenation catalyst 4 and 6, respectively. Numerals 8 and 10 designate cyclones positioned in the upper part of reactors 1 and 2, repectively, 'and numerals 12 and 14 designate conduits connected to the lower part of cyclones 8 and 10 for returning separated powdered catalyst to the main catalyst bodies 4 and 6, respectively.

Numeral 16 designates a conduit for conveying a mixture of oil and hydrogen charge through heater 18, valve 20 and thence into the base of .reactor 1 by way of conduit 22. Numeral 24 designates a conduit for conveying oil and hydrogen charge into the base of reac-tor 2 by way of valve 26 and conduit 28. The ends of conduits -22 and. 28 which terminate inside reactors 1 and 2, respectively, .are provided with a conventional distributing element (not shown) to introduce the charged gases, liquids and vapors evenly into the mass of catalyst particles in the base of the reactor. Numeral 30 designates a conduit for conveying hydrogen through heater 18 and thence into the base of reactor 1 or 2 by way of conduits 22 or 28, respectively, as determined by the setting of valves 32 and 34. Numeral 36 designates a conduit for delivering regeneration gas to conduit 38 at a rate controlled by valve 37 and numeral 39 indicates a conduit for' introducing purging steam at a rate controlled by valve 41. The setting of valves 40 `and 42 permits the ow at suitable times of regeneration gas or steam into reactor 1 or 2.

The exhaust side of cyclones 8 and 10 are connected, respectively, to conduits 44 and 46, which conduits are interconnected by conduit 48 provided with a withdrawal lead S4 which conveys the reaction products to a suitable system for recovery as determined by the setting of Valves 50 and 52. Numeral 60 designates a conduit for removing flue gas during regeneration from conduit 44 at a rate controlled by valve 62. Numeral 64 designates a similar conduit'for withdrawing flue gas from reactor 2 at a rate controlled by valve 66. Numerals 68 and 7 2 designate pumps for introducing the ui'dsin conduits 16 and 30, respectively.

Y rIt is desirable to employ a system which will handle constant ow of charge stock so that the advantage of continuous operation can be obtained. This can be yacoomplished by regenerating one of the reactors while the other reactor is on-stream and having regeneration completed and the regenerated reactor at reaction temperature so that the on-stream flow can be switched from one reactor .at the end of the reaction period to the regenerated reactor. yIt is convenient to operate so that the on-stream period in one reactor is equal to the period required for hydrogen purging, steam or inert g-as purging, regenerating and again steam or insert gas purging the other reactor.

Assuming that reactor 1 is on-stream and reactor 2 has been taken ofi-streamand is to be purgedI with hydrogenv and regenerated, valves 26, 32, 37, 40, 41, 42, 62 and 66 will be closed and valves Ztl, 34, Sil and 52 will be open. Hot hydrogenv gas then owsthrough reactor 2 and purges the hydrocarbons out of the reactor and hydrogenates part ofv thecarbonaceous deposit into hydrocarbonsV andv thus reduces the `amount of carbon tov be regenerated.. The hydrogen purge products then flow through conduits 46, rand valve 52 into conduit54.4 This hydrogen purgecan be-omittedif. desired but this will result in a higher regenerationl requirement. After the: hydrogen purge-1s cornpleted valves 34 and 52 are closed and valves 41, 42 and 66 are opened to. purge. reactor 2 with steam or inert gas from line .3.9. Then valve. l1 is closed and valve 3.7 is

opened. vRegenerationgas such as a mixture of air and steam or inert gas then ows upwardly through the body of .powdered catalyst 6 in reactor 2 and by combustion removes. the carbonacecus deposit thereon. This upward flow of gas and vapor is at such a rate as to cause suspension and turbulence-,of the catalyst particles. The regeneration gases may be those conventionally used such as a mixture of steam and/ or flue gas with air, the air content being sufficient to give a suitable regeneration temperature preferably of 950 tol ll50 F. more or less. Thecombustion products andA a small amount of the catalyst. powder flow through cyclone 10j. The catalyst powder 1s separated and returned through l-ine 14 and the combustion products ow through line 46, Valve 66 and line 64 to a suitable disposal system for the flue gas. When regeneration is; completed reactor 2 is purged with steam and Cooled to reaction temperature. This purging with steam is accomplished by closing valve 37 and opening valvey 41. After steam purging is completed valves 41 and 42 are closed.

While reactor 2 is being regenerated as described, the hydrocarbon charge stock and hydrogen at approximately reaction pressure flow through pump 68, line. 16, heater 1,'8, valve 20 and line 22 int-,o the base of reactor 1. The temperature to which the hydrocarbon and hydrogen is heated in heater 16 is that which will result in the necessary heat bal-ance, i. e., the desired reaction temperature in the reactor.

Since the-reaction is exotherm'ic only a part of the heat need be furnished by the charged mixture. From the standpoint of economy it is desirable to avoid thermal cracking in the preheater as much as possible. In other words, thermal cracking of heavy portions in the preheater results in carbon formation whereas if those portions are introduced onto the catalystthey are converted into valuable hydrocarbon products with relatively' little carbon formation. The amount of vapor-ization in the preheater can be varied greatly and will, of course, depend on the character of the hydrocarbon treated. All or substantially all of the hydrocarbon may be in liquid phase when introduced into the reactor or on the other hand all' of the volatilizable components .and much of the difficultly vaporizable components can be vaporized. This last method of operation is of course less satisfactory but is included within the scope of our invention since it utilizes the advantages thereof, i'. e., that portion of the diicultly vaporizable component which remains in liquid phase as well as the vapors are economically converted to valuable hydrocarbons in accordance with our invention. The optimum method of operation is to preheat so as to introduce into the reactors all of the diflicultly vaporizable components in liquid phase and all of the vaporizable components (i. e., vaporizable under the conditions withoutl decomposition) in vapor phase. These diflicultly vaporizable components are preferably introduced onto the catalyst in liquid phase as fast as they are converted into volatile products, but notso fast as to form av 'bridge of liquid between the catalyst particles as explained 'hereinafter.

The mixture of hydrogen and hydrocarbon oil upon entering the base of reactor `1 is distributed evenly throughout the lower portion of the catalyst bed and ilowsupwardly therethrough at such at rate as to maintain the catalyst particles' in a suspended and turbulentstate. The portion of the charge which is in liquid condition is absorbed on the catalyst particles and because of the mo' tion of the particles will be rapidly carried away from the poi-nt of absorption. A constantly fresh stream of eatlalyst particles therefore will be presented to the liquid portion of the charge so that it is absorbed and remQved Y 4; as fast as charged. llt is presumed that this liquid portion remains on the catalyst particle until it has been hydrol cracked to lower boiling hydrocarbons and these lower boiling hydrocarbons are then vaporized and pass upwardly with the vapor portion of the charge stock. These vapors are also destructively hydrogenated. The reaction products pass through cyclone 8 with a small amount of the catalyst powder. This catalyst powder is removed in cyclone 8 and returned to the catalyst bed through conduit 12. The reaction products then flow through conduit 4,4, Valve 50 and conduit 54 into a suitable product recovery system (not shown).

`When catalyst in reactor 1 has accumulated sufficient carbon to reduce its activity so as'to require regeneration, the mixture of hydrocarbon and hydrogen charge is -shunted to reactor 2 and reactor 1 is then reactivated by the step of hydrogen purging, steam purging, regeneration and steam purging as described in connection with reactor 2.` Thus valveZil is closed and valve 32 is opened to permit hydrogen under pressure to pass from pump 72', through conduit 3ft,` heater 1S and' conduitZZ into reactor l where ay part of, the coke deposit on the catalyst is hydrogenated to hydrocarbons as previously described. At the: same time valves Z6- and 52 are opened and valve. 66 is closed. The series of reactivation steps described in detail in connection with react-or 2 are then carried out in reactor 1. At the same time the on-stream reaction is carried' out in reactor 2. The process is thus carriedl out continuously and alternately in the two reactors. It is good practice to-keep the catalyst in a turbulent and suspended condition at all times, including'v thoseV times when switchingfrom on-stream to regeneration. Therefore switch valves are; used so that uidization is maintained during switchingprocedure.

The catalyst particles, in the reactor must be maintained in a suspended and turbulent'. condition by the hydrocarbon vapors and hydrogen gas duringl the on-stream period.

or by the regeneration gases during the regeneration period. Any method which will result in this suspended and turbulent condition can be used. It is preferred. that the catalyst powder be in the form of a uidized dense bed,l such as is well known in the fluidized catalytic cracking art. However, it is possible to` utilize a bed in which the particles are either more closely packed or more dispersed and agitated than is the' dense phase of a conventional duid; catalytic crackingbed. All that is necessary isl that the catalyst particles be suspended and be free to move with respect tol each other. This condition is attained if sufficient gas and vapor are passed through the catalyst particles to expand the catalyst particle bed to the point where. the particles become disengaged from each other and have; a random motion. lf the amount of expansion is insufficient to cause the particles to have a random motion, the catalyst will not be in a suitable condition for the catalytic operations contemplated herein, nor for thatrnatter will it be suitable for conventional fluidized catalytic operations. However, if the amount of gas and vapor passing through the catalyst bed C0: is gradually increased, the bed gradually expands and reaches a poi-nt where the particles become suspended and free to move in the gas or vapor stream. This is the minimum condition under which iiuidization can be said to exist and is the minimum condition of turbulence necessary for our invention. Calcula-tion and experiments have indicated that this results in a dense iluidized turbulent bed in the lower portion of the reactor andV a light phase of catalyst particles in the upper portion of the reactor. However, this minimum condition of turbulence and suspension is generally considerablyless than is employed in a conventional fluid catalytic cracking bed.

VIgt is advantageous to operate in or near this minimum condition of turbulence. The linear velocities employed in maintaining the catalyst in this condition of low turbulence are lower than have heretofore beenl used in iluid catalytic cracking. It has been found that a linear velocity range of 0.01 to 0.5 ft./sec. has been suniciently great to satisfactorily maintain low turbulence ad vantageous operating-v condit-ions. (This compares with a velocity` of l to 2 feet per second usually employed in catalytic cracking.) These low linear velocities are also advantageous from the standpoint of catalyst attrition. Production of catalyst lines is a function of the magnitude ofthe linear velocity and since the velocity employed in our proces-s may be low, correspondingly lov-J attrition rate is. obtained.

If the volume of gas and vapor is further increased, the dense phase undergoes further expansion while becoming less dense and bubbles in a manner similar to a boiling liquid. This condition gives greater turbulence. It is a condition which approximates that usually encountered in conventional iiuid catalyst beds. However, it is advantageous to keep the degree of iluidization and turbulence to a minimum which will give the desired degree of conversion since the lower degree of uidization does not require as long a reactor, which is important for pressure reactions, and also does not need as large capacity cyclones to remove the fines from the effluent gases and vapors. A still further increase in volume of gas and vapor owing through the bed will result in actual transport of the catalyst particles with the vapor and gas. This condition is also satisfactory but it is necessary that the length of the reactor be increased to allow sufficient contact time. If such a transport operation is to be used, it would be best to add the catalyst to the vapor-gas stream,

transport the mixture through the reactor and then sepaf rate the vapors and gas from the catalyst particles and return them for reuse.

The operation may be carried out without introduction or removal of catalyst to or from the reactor during the on-stream period, i. e., with the catalyst maintained as a fixed iluidized bed. In such event regeneration is effected either by regenerating the catalyst in situ, i. e., in the reactor, at intervals between on-stream periods, or by transferring the catalyst from the reactor to a separate regenerator at the end of the on-stream period, and replacing it with fresh or regenerated catalyst.

Where the velocity of the mixed hydrocarbon vapors and hydrogen is such that substantial amounts of the catalyst are carried out of the reactor during the onstream period, it is ordinarily desirable to maintain a relatively uniform amount of catalyst in the reactor during the on-stream period by introducing fresh, regenerated, or recirculated catalyst to the reactor during the on-stream period to replace that removed in the vapors leaving the reactor. In such case catalyst can be separated from the vapors leaving the reactor by means of suitable devices such as screens or a Acyclone separator from which the collected catalyst can be returned directly to the reactor r to a regenerator.

In practicing our invention, moving bed fluidized operations, such as are ordinarily used in uidized catalytic cracking, may be used. In such an operation the catalyst would be continuously withdrawn from the reactor, continuously introduced into a regenerator at approximately the same pressure as the reactor and continuously returned to the reactor from the regenerator. 'Ihis system has the advantage that intermittent shutdown for the regeneration is unnecessary. The removal of the catalyst from the reactor to the regenerator and the removal of the catalyst from the regenerator to the reactor can be accomplished by taking advantage of the catastatic or manometric principle utilized for moving iiuidized catalyst in conventional catalytic cracking units. However, this type of operation requires careful temperature control since the effect of the relatively high pressure conditions existing during our invention can easily upset the pressure differentials required for this method of circulating catalyst.

Regeneration at atmospheric pressure is desirable in certain cases. The apparatus illustrated in the drawing would of course be satisfactory for such regeneration. However, to operate continuously with a lluidized moving bed as described in the preceding paragraph and at the same time regenerate at atmospheric pressure, would require a method for withdrawing catalyst, regenerating it at atmospheric pressure and reintroducing it into the high pressure system. This can be accomplished by the continuous withdrawal of catalyst particles, passing them through a depressuring chamber, regenerating them at atmospheric pressure, cooling the catalyst particles and introducing the catalyst particles into the hydrocarbon charge before it passes into the reactor. Due to the fact that a low catalyst-to-oil ratio can be employed in our invention, it is feasible to introduce the catalyst into the hydrocarbon charge in this manner. Thus, destructive hydrogenation is an exothermic reaction, and it is unnecessary to employ the catalyst as a source of heat as in uidized catalytic cracking operations. Therefore much smaller amounts of catalyst can be used than are necessary in ordinary catalytic cracking where the reaction is endothermic and relatively huge amounts of hot catalyst are necessary to furnish the heat of reaction. When a lluidized catalyst has heretofore been used as a source of heat in a catalytic cracking operation it has been necessary to use it in catalyst-to-oil ratios of between about 5:1 and 30:1. However, in accordance with our invention it is in general quite satisfactory to use a much lower catalyst-to-oil weight ratio such as preferably between about 1:2 and 1:16 although lower or higher ratios can be used.

The degree and depth of conversion is largely dependent upon the reaction temperature and the space velocity. space velocity of between about 0.1 and 5 unit weights of hydrocarbon per unit weight of catalyst per hour gives a satisfactory degree of conversion at a temperature between about 750950 F. However, it is possible to employ lower or higher space velocities and obtain greater or lesser conversion and our invention therefore is not limited to this space velocity range. A lower space velocity will in general be best for oils containing the largest amounts of diicultly vaporizable materials, while the higher space velocities will be used for oils containing small amounts of these materials. The higher the temperature the greater will be the conversion. However, temperatures above about 950 F. result in excessive conversion to gas and coke. Temperatures below about 750 F. give insuicient conversion.

The process of our invention is applicable to any type of charge stock which contains difiicultly vaporizable hydrocarbons, by which term we mean hydrocarbons which cannot be vaporized in conventional commercial heaters at the reaction pressure without substantial decomposition. Examples of such materials are crude petroleum, reduced crude, topped crude, shale oil and heavy residual hydrocarbon oils. Ordinarily, introduction of these materials in liquid phase onto a catalyst at the temperature contemplated herein would result in a rapid deposit of a large amount of carbon or coke which would necessitate discontinuance of the operation. However, we have found that this does not take place under the conditions and manner of operation described herein. The liquid portion of the charge is absorbed on the catalyst granules where it becomes cracked and/or hydrogenated with formation of lighter products and with very little carbon formation. Not only does the step of charging partly in liquid phase result in less carbon production but it also has an added advantage Yin that the liquid hydrocarbon absorbed on the catalyst acts as a hydrogen donor thus increasing the rate of hydrogenation. Instead of getting a large amount of carbon and gas, as would be expected, a high degree of conversion to high grade liquid products is obtained. This hydrocarboncontent is also advantageous since the conversion of these heavy materials is slow and they are retained on the vcatalyst until converted. It is therefore evident that we operate with a catalyst which at all times has a high hydrocarbon content. Analysis indicates that the catalyst absorbs and holds from about 20 to 50 per cent by weight of hydrocarbon during the'reaction. Carbon or coke is also present and is gradually built up in addition to this.

Our invention is of particular value in connection with the treatment of low-grade heavy hydrocarbon mixtures such as those which contain large amounts of asphalt or sulfur or those which have a high carbon residue. The asphalt and high carbon residue components are largely converted into relatively low boiling hydrocarbons of good quality and the sulfur components are largely converted into hydrocarbons and hydrogen sulfide which can be separated with relative ease from the product. This desulfurization is of the catalytic type. Thus the sulfur compounds are catalytically hydrogenated to hydrogen sulfide which is continuously removed with the product and must be separated therefrom. Some sulfur may be absorbed on the catalyst when certain metals or oxides are used, but only in minor amount.

While a temperature range of about 750 to 950 F. may be used, we prefer to utilize a temperature between about 800 and 875 F. for most heavy charge stocks since this temperature range gives a high conversion to products within the gasoline boiling point range and further reduces the small amount of coke that is formed. Ordinarily, coke formation increases with temperature increase if other conditions are equal. However, under the conditions of our invention coke formation reaches a armonia minimum between 800 and `8775" F., i. e., :it is :higher below and y"above this temperature range. This yis there- 'fore 'an especially :advantageous temperature range to use. The pressure may be between -about 250 and 2000 p. s. i. However, we prefer va pressure between about 500 and 1000 p. -s. i. This latter pressure lrange is of definite advantage :since it reduces the acost .of 'the apparatus and simplifies the operation of the process without sacrifice :in yields or quality. The fact that pressures below 2000 p. s. i. can be employed without voxcessive carbon formation is :quite contrary to expectation, especially when one considers that a liquid Vhydrocarbon is present. At the same time these ,low pressures :give quite satisfactory results. A hydrogen-tooil ratio of between 300 and 20,000 cubic feet of hydrogen per barrel lof liquid oil charged may be used. A ratio of between about 5000 and 10,000 is preferred.

The degree of vaporization lof the hydrocarbon vcharge prior to introducing into the reactor must not lbe such that constituents in Aliquid phase are introduced at a rate t suihcient to cause them to accumulate in the reactor to yform a `body of Aliquid of substantial size, or at such a rate as to result in excessive wetting of the catalyst particles since this would prevent the lnecessary particle suspension or turbulence. It is advantageous to so control the vaporization that the total liquid yin the reactor is always less than that which can be absorbed on the catalyst. If the catalyst kparticles accumulate enough liquid to form a bridge between the particles, .agglomeration will result and the suspended, turbulent condition will terminate or become imperfect. Suspension and turbulence of the catalyst .is necessary in order to get eicient contact between the .catalyst particles and the reactants, which contact is necessary for eective catalysis of the reaction. It is also necessary so that eticient heat exchange between particles will take place. The reaction is exothermic and if eiiicient intermixing of catalyst particles does not take place, hot spots will develop where the reaction is going on at an excessive rate and temperature. This will result in 'excessive conversion to gas and coke, deactivation of the catalyst and perhaps damage to the reactor. The degree of turbulence or interi-nixing can be regulated by controlling the ratio of the charged vapors and hydrogen to the catalyst.

Any nely divided hydrogenation catalyst can be employed. Examples of suitable catalysts are molybdenum, tungsten, vanadium, chromium, cobalt, nickel, iron and tin and their oxides and sulfdes. Mixtures of these materials or compounds of two or more of these oxides are advantageously employed. compounds of the iron group metal oxides or suldes with the oxides or suldes of Group VT left column of the periodic table constitute very satisfactory catalysts. Examples of such mixtures or compounds are nickel molybdate, tungstate or chromate (or thio molybdate,

thio tungstate or thio chromate) or mixtures of nickel oxide with molybdenum, tungsten or chromium oxides. These catalysts are advantageously deposited upon or otherwise composited with a porous carrier such as activated alumina, silica gel or the various synthetic or natural silica-alumina type cracking catalysts -or other refractory materials having a large surface area. The composite of hydrogenating catalyst and carrier is prepared in known manner such as by impregnating the carrier particles with a solution of a compound or salt of the desired hydrogenating component followed by calcining (and reduction if a reduced catalyst is to be used). While we prefer to employ porous carriers, non-porous carriers may be employed as also may powders composed entirely of the hydrogenating component.

The size of the catalyst particles ycan vary considerably, the only requirement being that it should be small enough to be 'suspended by the Vcurrent of gas and vapor passed through the catalyst bed. However, it is necessary to employ catalyst particles which are not so small as to be largely carried along by the gas and vapor stream, if a non-transport type of operation is employed. However, this danger of losing the catalyst or the finer part of the catalyst is not as great under the pressures employed herein as is the case under the approximate atmospheric pressures usually employed for catalytic kcracking. Particles having a diameter falling between about 400 and 50 meshare satisfactory. Most commercial catalysts are a mixture of particles having a variety of diameters but these particles are almost entirely within the For instance, mixtures or .f5.1

E8 above diameter range. It is advantageous 'to use such mixtures but those containing .large amounts lof large or ne material, Yi. e., near :or larger than 50 imesh or .near or smaller .than 400 mesh,.should `be avoided.

It is advantageous to recycle the :hydrogen separated from thereaction products for reuse in the process. During the operation of the process hydrocarbon gases such as methane, ethane, propane, etc. are formed and contaminate the hydrogen. However, we have found .that recycle hydrogen containing large .amounts lof such yhydrocarbon Igases is yquite satisfactory. -It is also evident that .impure hydrogen can be used when the recycling step is not'used.

EXAMPLE I A hydrocracking catalyst vconsisting of 1.5 per 'cent nickel oxide and 4.71 per cent Atungsten oxide deposited on microspheroids of a synthetic silica-alumina vcracking base was charged to a reactor. The lcatalyst had 'the fol lowing Roller analysis:

Approx Micron Range Wtc'cter' Mesh Size The catalyst was reduced by passing yhydrogen through the reactor at atmospheric pressure and 900 F. for ltWo hours. The catalyst was then cooled to 800 F. The reactor was then pressured with hydrogen to 500 p. s. i. Kuwait crude and hydrogen were heated to 800 in a preheating coil and charged to the reactor which was maintained at 800 F. Analysis of the Kuwait crude is contained in Table Ill. The hydrogen stream was 99.9 per cen-t pure. The hydrogen was 16.3 weight per cent of the oil charged. The percentage of hydrocarbon in the liquid phase was approximately 10 per cent. During the on-stream period '3.48 weights `of loil per unit'weight of the catalyst were charged to the reactor at a space velocity of 0.87 (wt. oil/hr./wt. cat). The oil feed was then stopped, The reactor was purged with hydrogen at a ilow rate of 85 per cent of the onstream rate while at the same temperature and pressure. After hydrogen purging, the unit was depressured, purged with atmospheric nitrogen and regenerated with preheated air. During the regeneration period a maximum temperature of 1023 F. was reached during the rst hour. The regenerated catalyst then was cooled to about l800 F. and anew cycle was started, i. e., the reactor was pressured anda heated mixture of Kuwait crude and hydrogen were again contacted therewith. The `length L of Vthe on-stream period was approximately 4 hours,

the hydrogen purge period was approximately 3 hours and the regeneration period was about 5 hours. Table I gives the details on the analysis of the hydrocarbons produced during the on-stream and hydrogen purging periods.

Table I Recovery (weight per cent of charge): Gas (C1-C3) 6.9 Liquid product 90.5 Sulfur 2.2 Carbon 2.5

Total 102.1

Inspection data:

Total product- Gravity-A. P. I 46.4 Sp. Gr 0.7954 Sulfur (weight percent) 0.21 DistillationyPercent at 392 53.0 Percent at 500 70.2 Percent at 5-90 c 83.5

9 EXAMPLE II Hydrocracking catalyst particles consisting of 1.0 per cent NiO and 3.3 per cent W03 on activated alumina stabilized with silica were charged to a reactor as described in Example I. However, the catalyst was not reduced with hydrogen as in Example I. The reactor was pressured with nitrogen at a pressure of 500 p. s. i. and a temperature of about 840 F. was established in the reactor while the flow controller was calibrated with hydrogen ow which yby-passed the reactor case. Hydrogen at the rate of 10,000 cu. ft. per barrel and Kuwait crude (Table III) were heated to 830 F. in a preheating coil and then charged to the pressured reactor which was maintained at 840 F. The hydrogen stream was 99.9 percent pure. During the on-stream period 3.38 weights of oil per unit weight of the catalyst were charged to the reactor at a space velocity of 0.85 (wt. oil/hr./wt. cat). the reactor was purged with hydrogen at the reaction temperature and pressure at a hydrogen flow rate of 96 percent of the on-stream rate. After hydrogen purging the unit was depressured, purged with nitrogen and regenerated with preheated air. The regenerated catalyst then was cooled to reaction temperature and a new cycle was started. The length of the on-stream period was approximately 4 hours and the hydrogen purge period about 2 hours. Liquid product was collected during the on-stream and hydrogen purge periods. The product had the characteristics given in Table II.

Table II Recovery (weight per cent of charge) Gas (C1-C3) 4.8 Liquid product 88.4 Sulfur 2.4 Carbon 1 3.6

Total 99.2

Inspection data:

, Total product- Gravity- API 45.0 Sp. Gr 0.8017 Sulfur (weight per cent) 0.128 Distillation- Per cent at 392 50.0 Per cent at 500 72.0 Per cent at 590 86.4

lCarbon from analysis of regeneration gas.

Table III INSPECTION DATA OF' KUWAIT CRUDE Gravity:

API 31.4. Sp. Gr 0.8686. Viscosity:

100 F., centistokes 9.96.

SUS 58.7. Color 8-|-dilute. Pour point, F Flash (P. M.), F 75 F. Carbon residue, per cent on material above 590 F 8.53. Sulfur, per cent (bomb) 2.49.

Salt, lbs./ 1000 bbl 5.2

Distillation:

Per cent at 392 F 24.0. Per cent at 500 F 34.8. Per cent at 590 F 43.5. Recovery, per cent 43.5. Residue, per cent 53.5. Loss, per cent 3.0.

EXAMPLE III Hydrocracking catalyst particles composed of per cent nickel tungstate deposited on microspheroids of a .synthetic silica-alumina cracking catalyst base were charged to a reactor as described in Example II. The reactor was pressured with nitrogen at a pressure of 1000 p. s. i. and a temperature of about 850 F. was established and maintained in the reactor. Hydrogen at the rate of 9000 cu. ft. per barrel and Baxterville crude (Table V) were heated The flow was then stopped and proximately 2 hours.

to'840 F. in a preheating coil and then charged to the pressured reactor. During the on-stream period 3.85 unit weights of oil per unit weight of the catalyst were charged to the reactor at a space velocity of .47 (wt. oil/wt. cat./hr.). Thev flow was then stopped and the reactor was purged with hydrogen at the reaction temperature and pressure at a hydrogen How rate of approximately the on-stream rate. After hydrogen purging, the unit was depressured, purged with an inert gas and regenerated with preheated air. The regenerated catalyst then was cooled to approximately 850 F. and a new cycle was started. The length of the on-stream period was approximately 8 hours, and the hydrogen purge period ap- Liquid product was collected during the one-stream and hydrogen purge periods. The product had the characteristics given in Table IV.

Table IV Recovery (weight per cent of charge) Gas (C1-C3) 10.2 Liquid product 78.8 Sulfur 2.1 Carbon 7.0

Total 98.1

Inspection data:

Total product- Gravity- API 44.4 Sp. gr 0.8044 Sulfur (weight per cent) 0.36 Distillation- Per cent at 392 F 54.2 Per cent at 500 F 72.7 Per cent at 590 F 87.2

Table V INSPECTION DATA OF BAXTERVILLE CRUDE Gravity:

API 15.8. Sp. gr 0.9609. Viscosity:

F. centistokes 608. SUS 2811. Color". Black (8+ dilute). Pour point Below 0 F. Flash (P. M.), F 120.

Carbon residue, per cent on material While the present invention resembles in certain respects conventional destructive hydrogenation, it diers therefrom in that it involves the utilization of lower pressure than ordinarily used therein. This lower pressure results 1n considerably less hydrogen consumption. This reduc'- tionV in amount of hydrogen consumed is of unusual importance because of the high cost of hydrogen. As a matter of. fact this is one of the reasons destructive hydrogenation has not been widely used in the refining of petroleum and like hydrocarbons. In the runs described in the Examples the hydrogen consumption varied from about 70 to 800 cu. ft. lof hydrogen per barrel of oil. This corresponds to about 0.17 to 1.34 weight per cent. As compared with this the hydrogen consumption during desructive hydrogenation of analogous hydrocarbons involved hydrogen consumption of approximately 4 to 5 weight per cent.

It is known that as the pressure in a hydrogenation process is increased, other conditions being equal, the tendency toward gasification of the hydrocarbon charge is encouraged. 'The fact that we can use lower pressures than ordinarily used for destructive hydrogenation therefore results in decreased tendency toward high gasification. In addition the low temperature gradient over the entire fluid bed, i. e., more near-ly constant temperature throughout the uid bed, presents a condition which gives greater conversions with correspondingly low coke deposits. This is accomplished under pressures which would lead to excessive coke laydown in lixed granular or pellet catalyst beds where temperature gradients of .greater magnitude are obtained.

As .previously indicated, the invention is unique in eiecting conversion of heavy stocks into lower boiling products with relatively little formation of carbon and in very high yields, as compared with conventional coking or catalytic cracking methods (i. e., those methods other than destructive hydrogenation, usually employed to convert heavy hydrocarbons). vThis is, to a considerable extent due to the absorption of the liquid and avoidance of accumulation of a separate liquid phase in the reactor which would cause rapid coking and formation of a slurry which would impair fluidization. This liquid absorption decreases carbon formation by increasing the rate at which the liquid is hydrogenated and this in turn results in an increase in cracking rate. Also, this absorption of liquid has another advantage since it provides a means for keeping the liquid oil and catalyst together until cracking -is complete. In other words, the liquid portion, which requires a longer time to hydrogenate and crack, remains on the catalyst a preferentially longer time.

What we claim is:

1. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains difliculty vaporizable hydrocarbon components in order to form `lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it contains a substantial amount of the diicultly vaporizable -components in liquid phase and hydrogen into contact with particles of Va porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 750 and 950 F. and at a pressure between about 250 and 2000 p. s. i. g., controlling the .rate of introduction of the hydrocarbon mix'- ture so that a substantial amount of the liquid portion of the diicultly vaporizable hydrocarbon components `of the mixture is at all times absorbed in the pores Aof the catalyst particles -but so that the total amount :of liquid in contact with the catalyst particles at any time is vless than the total amount which-can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a fluidized condition i. e., turbulent and suspended in the hydrogen Yand hydrocarbon vapors, and separating catalyst particles from the hydrogen and the hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming suspended in a liquid body of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as well as the hydrocarbon vapors are eiciently converted by destructive hydrogenation into lower boiling hydrocarbons.

2. In a process for the catalytic destructive hydrogenation of. a hydrocarbon mixture which contains diicultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it contains a substantial amount of the diicultly vaporizable components in liquid phase and hydrogen into contact with particles of a porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 800 and 875 F. and at a pressure between about 500 and 1000 p. s. i. g., controlling the rate of introduction of the hydrocarbon mixture so that a substantial amount of the liquid lportion of the diiicult-ly vaporizable hydrocarbon components of the mixture is at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any ytime is less than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a uidizcd condition, i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors, `and separating catalyst particles 'from the hydrogen and the hydrocarbon .reaction products, whereby Vthe catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming .suspended in a liquid Vbody of substantial size and the liquid hydrocai-bons absorbed in the pores of the 'catalyst particles as well as the Vhydrocarbon vapors are ei'ciently converted Civ "iti

by destructive hydrogenation into lower boiling hydrocarbons. v

3. In a process for the catalytic destructive hydrogenation 'of a crude petroleum which contains dicultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said crude .petroleum while it contains a substantial amount of the diiicultly vaporizable components in lliquid phase and hydrogen into contact with particles of a porous hydrogenation catalyst which are suspended -i-n hydrogen and hydrocarbon vapor, at a temperature between about 800 and 875 F. and at a pressure between about 500 and 1000 p. s. i. g., controlling the rate of introduction of the crude petroleum so that a substantial amount of the liquid portion of the diiicultly vaporizable hydrocarbon components is at all times absorbed in the pores of the catalyst particles but so that the tota-l amount of liquid in contact with the catalyst Vparticles at any time is less than the total amount which can be 'absorbed in the pores of the catalyst particles, maintaining 'the 'catalyst particles in a fluidized condition, i. e., turbulent and suspended in the hydrogen and crude petroleum vapors, and separating catalyst particles from the hydrogen and the hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming suspended in a liquid body o'f substantial size and the liquid hydrocarbons absorbed in the pores of the Acatalyst particles as well as the vapors are eiciently converted by destructive hydrogenation into lower boiling hydrocarbons.

4. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains diicultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it contains a substantial amount of the difficulty vaporizable components in liquid phase and hydrogen into contact with particles of a porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor and which are selected from the group consisting of iron group metal oxides and suliides in combination with a member of the group consisting of oxides and sultdes of metals of Group Vl left-hand column of the Periodic Table composited with a porous carrier, at a temperature between about 750 and 950 F. and at a pressure between about 250 and 2000 p. s. i. g., controlling the rate of introduction of the hydrocarbon mixture so that a substantial amount of the liquid portion of the diflicultly vaporizable hydrocarbon components of the mixture is at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any time is less than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a fluidized condition, i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors, .and separating .catalyst particles from the hydrogen and the .hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming substantially .agglomerated and are prevented from becoming .suspended in a liquid body of substantial size `and the liquid .hydrocarbons absorbed .in the pores of the -catalyst .particles Vas well as .the hydrocarbon vapors are eiciently converted by destructive hydrogenation into lower boiling hydrocarbons.

5. ln a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains difficulty vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it `'contains a substantial amount of the diticulty vaporizable components in liquid phase and hydrogen into `contact with -a -fixed bed of particles of a porous hydrogenation catalyst which particles are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 7 50 and 950 F. and at a pressure between about 250 and 2000 p. s. i. g., controlling the rate of introduction of the hydrocarbon mixture so that a substantial amount ofthe liquid portion .of the dicultly vaporizable hydrocarbon components of the mixture is at all times absorbed in `the pores of the catalyst particles but so that the total amount of .liquid in contact with the catalyst particles 'at any time i's less than the total amount which can be 'absorbed in the y'pores of the catalyst particles, maintaining the catalyst particles in a 'iluidized condition,

i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors, separating catalyst particles from the hydrogen and the hydrocarbon reaction products, continuing said treatment with the same catalyst particles until the catalyst particles require regeneration, terminating the treatment, regenerating the catalyst particles and re-using the regenerated catalyst particles in the process, whereby the catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming suspended in a liquid body of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as well as the hydrocarbon vapors are eiliciently converted by destructive hydrogenation into lower boiling hydrocarbons.

6. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains ditlculty vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it contains a substantial amount of the ditlicultly vaporizable components in liquid phase and hydrogen at a linear velocity of between about 0.01 and 0.5 foot per second into contact with particles of a porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 750 and 950 F. and at a pressure between about 250 and 2000 p. s. i. g., controlling the rate of introduction of the hydrocarbon mixture so that a substantial amount of the liquid portion of the diflicultly vaporizable hydrocarbon components of the mixture is at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any time is less than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a uidized condition, i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors and separating catalyst particles from the hydrogen and the hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming suspended in a liquid body of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as well as the hydrocarbon vapors are eiciently converted by destructive hydrogenation into lower boiling hydrocarbons.

7. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains diicultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises preheating said hydrocarbon mixture to form a liquid-vapor mixture without decomposing a substantial amount of the dicultly vaporizable components, introducing this liquid-vapor mixture and hydrogen into contact with particles of a porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 750 and 950 F. and at a pressure between about 250 and 2000 p. s. i. g., controlling the rate of introduction of the hydrocarbon mixture so that between about 20 and 50 per cent by weight of liquid diicultly vaporizable components are at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any time is less than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a lluidized condition, i. e, turbulent and suspended in the hydrogen and hydrocarbon vapors, and separating catalyst particles from the hydrogen and the hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming suspended in a liquid body of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as well as the hydrocarbon vapors are eciently converted by destructive hydrogenation into lower boiling hydrocarbons.

8. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains difcultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it contains a substantial amount of the diflicultly vaporizable components in liquid phase and hydrogen into contact with particles of a porous hydrogenation catalyst which are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 750 and 950 F. at a pressure between about 250 and 2000 p. s. i. g., and at a space velocity between about 0.1 and 5, controlling the rate of introduction of thehydrocarbon mixture so that between about 20 and 50 per cent by weight of the liquid diticultly vaporizable components are at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any time is les's than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a lluidized condition, i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors, and separating catalyst particles from the hydrogen and the hydrocarbon reaction products, whereby the catalyst particles are prevented from becoming substantially agglomerated and are prevented from becoming suspended in a liquid body of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as Well as the hydrocarbon vapors are eflciently converted by destructive hydrogenation into lower boiling hydrocarbons.

9. In a process for the catalytic destructive hydrogenation of a hydrocarbon mixture which contains difficultly vaporizable hydrocarbon components in order to form lower boiling hydrocarbons therefrom, the improvement which comprises introducing said hydrocarbon mixture while it-contains a substantial amount of the difficultly vaporizable components in liquid phase and hydrogen in amounts of between 200 and 20,000 cubic feet per barrel, into contact with particles of a porous hydrogenation catalyst which particles are suspended in hydrogen and hydrocarbon vapor, at a temperature between about 750 and 950 F. at a pressure between about 250 and 2000 p. s. i. g., and at a space velocity between about 0.1 and 5, controlling the rate of introduction of the hydrocarbon mixture so that between about 20 and 50 per cent by weight of the liquid diicultly vaporizable components are at all times absorbed in the pores of the catalyst particles but so that the total amount of liquid in contact with the catalyst particles at any time is less than the total amount which can be absorbed in the pores of the catalyst particles, maintaining the catalyst particles in a fluidized condition, i. e., turbulent and suspended in the hydrogen and hydrocarbon vapors, separating catalyst particles from the hydrogen and the hydrocarbon reaction products, continuing said process until between about 1 and 16 unit weights of hydrocarbon mixture have been contacted with one unit weight of the catalyst particles, terminating said contacting, regenerating the catalyst particles and re-using the regenerated catalyst particles in the process, whereby the catalyst particles are prevented from becoming substantially agglomerated and are -prevented from becoming suspended in a liquid body/of substantial size and the liquid hydrocarbons absorbed in the pores of the catalyst particles as well as the hydrocarbon vapors are efficiently converted by destructive hydrogenation into lower boiling hydrocarbons.

References Cited in the file of this patent UNITED STATES PATENTS 1,972,948 Payne Sept. 1l, 1934 2,100,353 Pier et al Nov. 30, 1937 2,268,187 Churchill Dec. 30, 1941 2,517,900 Loy Aug. 8, 1950 

1. IN A PROCESS FOR THE CATALYTIC DESTRUCTIVE HYDROGENATION OF A HYDROCARBON MIXTURE WHICH CONTAINS DIFFICULTY VAPORIZABLE HYDROCARBON COMPONENTS IN ORDER TO FORM LOWER BOILING HYDROCARBONS THEREFROM THE IMPROVEMENT WHICH COMPRISES INTRODUCING SAID HYDROCARBON MIXTURE WHILE IT CONTAINS A SUBSTANTIAL AMOUNT OF THE DIFFICULTY VAPORIZABLE COMPONENTS IN LIQUID PHASE AND HYDROGEN INTO CONTACT WITH PARTICLES OF A POROUS HYDROGENATION CATALYST WHICH ARE SUSPENDED IN HYDROGEN AND HYDROCARBON VAPOR, AT A TEMPERATURE BETWEEN ABOUT 750* AND 950* F. AND AT A PRESSURE BETWEEN ABOUT 250 AND 2000 P.S.I.G., CONTROLLING THE RATE OF INTRODUCTION OF THE HYDROCARBON MIXTURE SO THAT A SUBSTANTIAL AMOUNT OF THE LIQUID PORTION OF THE DIFFICULTY VAPORIZABLE HYEROCARBON COMPONENTS OF THE MIXTURE IS AT ALL TIMES ABSORBED IN THE PORES OF THE CATALYST PARTICLES BUT SO THAT THE TOTAL AMOUNT OF LIQUID IN CONTACT WITH THE CATALYST PARTICLES AT ANY TIME IS LESS THAN THE TOTAL AMOUT WHICH CAN BE ABSORBED IN THE PORES OF THE CATALYST PARTICLES, MAINTAINING THE CATALYST PARTICLES 